Dehydrogenation of alkanes to unsaturated hydrocarbons, mainly to mono-olefins, is described in the literature and is practised on a commercial scale.
According to one proposal an alkane feed is contacted at sub-atmospheric pressure and at an elevated temperature with a preheated charge of catalyst, such as chromium oxide on alumina, in a fixed bed reactor. For further details of the catalyst reference should be made to U.S. Pat. No. 3,711,569. Due to the endothermic nature of the dehydrogenation reaction the catalyst is rapidly cooled on contact with the alkane feed. Carbon is deposited on the catalyst as the reaction proceeds. In order to provide the necessary heat of reaction it is usual after a short while to switch the alkane feed to another reactor whilst the catalyst of the first mentioned reactor is regenerated by burning off the deposited carbon with hot air. The heat liberated raises the temperature of the catalyst back to the desired level (e.g. about 640.degree. C.), whereupon further alkane feed can be supplied to the reactor. In a typical commercial plant there may be three such reactors, each of which remains on stream in turn for a short period (e.g. of the order of 7 to 10 minutes), before the catalyst has cooled to a temperature, e.g. about 540.degree. C., requiring reheating by burning off the accumulated carbon deposit.
This process has the advantage that little or no isomerisation of the product alkane occurs so that n-butane, for example, may be smoothly converted to a mixture of butene-1 and cis- and trans-butene-2, whilst iso-butane can be converted to iso-butylene without any significant amounts of n-butenes being formed. This means that product recovery is facilitated.
A disadvantage of this process is that it is a cyclic process which is subject to considerable temperature variation in operation. Due to its cyclic nature it is relatively complex to operate and the use of multiple reactors inevitably increases the capital cost. Moreover, since each cycle is very short the plant requires constant supervision and is expensive to operate in terms of labour costs. In addition this process is noted for its low selectively for olefin production and results in production of significant quantities of undesirable by-products. Another major disadvantage is that it is operated under vacuum and so the plant must incorporate not only vacuum equipment but also compression equipment which is required for product recovery.
Another proposal, which has proceeded as far as the pilot plant stage, is described in an article "Catalytic LPG dehydrogenation fits in 80's outlook" by Roy C. Berg et al at page 191 of Oil & Gas Journal for Nov. 10, 1980. According to this proposal a mixture of alkane and hydrogen is contacted with a platinum-containing catalyst in a number of series-connected stacked reactors at a temperature in the range of from about 550.degree. C. to about 600.degree. C. In this design a moving bed of catalyst is used in which catalyst is continuously withdrawn from the bottom of the reactor system and then passed to a regenerator in which it is continuously regenerated to remove carbon deposits and reheat the catalyst before being recycled to the top of the reactor system.
Although this proposal has the advantage of continuous reaction, isomerisation of product alkenes may occur. For example, it is estimated according to Table 4 of the above-mentioned article in Oil & Gas Journal that, in addition to 80 parts by weight of iso-butylene, there will be typically formed per 100 parts by weight of iso-butane feedstock 9 parts by weight of n-butenes. The separation of n- and iso-butenes is relatively difficult and so product recovery is complicated in this process. To maximise yield of iso-butene it is necessary to separate and recover the n-butenes, to hydrogenate these to n-butane, to isomerise this n-butane to iso-butane, and to recycle this to the hydrogenation process. Moreover the platinum-containing catalyst is susceptible to poisoning by impurities in the feedstock. Thus it is necessary to purify the feedstock rigorously in order to remove such impurities or at least to reduce their concentrations to acceptably low levels.
In yet another process (which, it is believed, has also not proceeded past the pilot plant stage) a mixed feed containing alkane and steam is contacted, in the absence of free oxygen, with a Group VIII metal catalyst supported on a highly calcined catalyst support such as alumina, silica or a Group II metal aluminate spinel. For further details regarding this process reference should be made to U.S. Pat. No. 3,641,182 as well as to U.S. Pat. Nos. 3,670,044; 3,692,701; 3,674,706; 4,005,985; 3,761,539; 3,957,688; 3,894,110; 3,880,776; 4,041,099; 4,191,846; 4,169,815; and 4,229,609. In this process a number of fixed tube reactors are used, the alkane feed stream being switched from one reactor to the other whilst the catalyst of the first-mentioned reactor is regenerated, typically by passing a mixture of steam and air through the catalyst.
Although this proposal has the advantage that the catalyst can be used for quite long periods between regenerations, e.g. several hours or so, it still suffers from the drawback of being a cyclic process and requires high capital investment.
There is accordingly a need to provide a continuous process for dehydrogenation of alkanes in which yields of product olefin are maximized with essentially no co-isomerisation to other olefins.